A model for natural gas-air mixture combustion in fluidized beds is proposed,. 19 counting ... The reactor is equipped with a perforated plate distributor of 1.8. 62.
* Revised Manuscript
MODELLING AND EXPERIMENTAL VALIDATION OF A FLUIDIZED BED REACTOR FREEBOARD REGION: APPLICATION TO NATURAL GAS COMBUSTION.
1 2 3 4 5 6 7 8 9 10
S. DOUNIT , M. HEMATI *, R. ANDREUX Laboratoire de Génie Chimique, U.M.R 5503, ENSIACET, 05 rue Paulin Talabot, 31106 Toulouse - France. ABSTRACT
11
A theoretical and experimental study of natural gas-air mixture combustion in a
12
fluidized bed of sand particles is presented. The operating temperatures are lower than a
13
critical temperature of 800 °C above which the combustion occurs in the vicinity of the
14
fluidized bed. Our study focusses on the freeboard zone where most of the methane
15
combustion takes place at such temperatures. Experimental results show the essential role of
16
the projection zone in determining the global thermal efficiency of the reactor. The dense bed
17
temperature, the fluidizing velocity and the mean particle diameter significantly affect the
18
thermal behaviours.
19
A model for natural gas-air mixture combustion in fluidized beds is proposed,
20
counting for interactions between dense and dilute regions of the reactor [Pré et al. (1998)]
21
supplemented with the freeboard region modelling of Kunii-Levenspiel (1990). Thermal
22
exchanges due to the convection between gas and particles, and due to the conduction and
23
radiation phenomena between the gas-particle suspension and the reactor walls are counted.
24
The kinetic scheme for the methane conversion is that proposed by Dryer and Glassman
25
(1973). Model predictions are in good agreement with the measurements.
26
Keywords: Fluidization, Combustion, Natural gas, Model, Freeboard
27
I: INTRODUCTION.
28
Fluidized bed combustors have many advantages, including their simplicity of
29
construction, their flexibility in accepting solid, liquid or gaseous fuel, and their high
30
combustion efficiency at low temperatures what minimises NOX generation [Foka et al.
31
(1994)]. Fluidized bed natural gas combustors are used in many industrial applications such as
32
incineration of sludge with high moisture content or solid particles calcination. Given the
33
ecological benefit of using natural gas in fluidized bed furnaces, it is of interest to describe the
34
combustion process in fluidized beds using experimental and theoretical approaches.
35
Natural gas combustion mechanisms in fluidized beds have been widely investigated
36
at bed temperatures greater than a critical temperature close to 800 °C [Dounit et al. (2001),
37
Pré et al. (1998), Pré-Goubelle (1997), Van Der Vaart (1992), Yanata et al. (1975), Baskakov
38
and Makhorin (1975)], above which the methane conversion is fully realized in the dense
39
zone. Sadilov and Baskakov (1973) reported temperatures measured in the bubble eruption
40
zone above the bed surface greater than the theoretical flame temperature. At bed
41
temperatures lower than this critical temperature, the combustion of methane is mainly
42
realized in the freeboard zone. In such conditions, the methane bubble eruption in the lean
43
phase of the fluidized bed induces high exploding risks [Dounit et al. (2001)]. Thus, low
44
temperature reacting fluidized bed applications require experiments for safety reasons. In such
45
a way, hydrodynamics and thermal phenomena occurring in the freeboard region of fluidized
46
bed reactors must be investigated and modelled.
47
However, there is a lack of references in the literature on the experimental and
48
theoretical study of the reacting fluidized bed dilute region. Indeed, previous works were
49
either room temperature hydrodynamic studies [Wen and Chen (1982), Fournol et al. (1973),
50
Zenz and Weil (1958), Kunii and Levenspiel (1990), George and Grace (1978)] or isothermal
51
catalytic chemical studies.
52
In this paper, hydrodynamics and thermal phenomena coupling in the projection zone
53
of a fluidized bed operating in the bubbling regime is experimentally studied and modelled.
54
The natural gas combustion modelling proposed by Pré et al. (1998) is used to predict the
55
reactor dense region while the Kunii-Levenspiel modelling (1990) is used to describe the
56
freeboard region. The two-stage kinetic scheme of methane conversion proposed by Dryer
57
and Glassman (1973) is used. The resulting reactor model is validated using our experimental
58
data.
59
II: DESCRIPTION OF THE SET-UP.
60
The experimental set-up is given in Figure 1. The reactor consists of a heat resistant of
61
steel pipe 180 mm in diameter and 1400 mm high, with a disengaging section 360 mm in
62
diameter and 1000 mm high. The reactor is equipped with a perforated plate distributor of 1.8
63
% porosity. The bed temperature is controlled using cooling air flowing in a double shell. The
64
pneumatic valve what controls the air circuit feeding into the double shell is controlled by a
65
PID system equipped with a thermocouple located in the dense bed. Natural gas with 97 %
66
methane content is used as the combustible and premixed to the air in the windbox. The
67
reactor is fitted axially with chromel-alumel-type temperature sensors and water-cooled tubes
68
to sample the gas located at 50, 100, 250, 300, 400, 450, 550, 600, 650, 700, 900, 1000, 1100,
69
1200 and 1300 mm above the distributor. The sampling tubes are connected to a pump, a
70
cooling unit to eliminate water, infrared-type analysers to measure CH4, CO2, and CO mole
71
fractions, and a paramagnetic-type analyser for O2 measurements. Experiments safety is
72
ensured using (i) a burner placed in the wider section of the reactor what converts the
73
unburned gas exiting at the fluidized bed outlet (ii) a thermocouple placed in the windbox
74
connected to a PID regulator what cuts off the reactor feeding when the windbox temperature
75
reaches a critical value chosen well below the air-methane auto-burning temperature.
76
The studied operating conditions are presented in Table 1. Typical sand particles with
77
a density of 2650 kg.m-3 and mean size ranging between 100 and 550 µm are fluidized. The
78
air factor, defined as the ratio of the volume of air fed into the reactor to the required volume
79
for stoichiometric combustion of methane, ranges between 1.0 and 1.5. The bed temperature,
80
Tbed, is defined to that measured 150 mm above the distributor. The fluidizing velocity ranges
81
between 2 and 4 times the minimum fluidizing velocity at 20 °C.
82
(Table 1).
83
III: EXPERIMENTAL RESULTS.
84
III.1: Typical experiment presentation.
85
C1 experiments (Table 1) are realized in permanent regime, with a dense bed
86
temperature ranging between 650 and 800 °C.
87
Temperature profiles.
88
The differences between the local temperature measured along the reactor axis and the
89
dense bed temperature, Tbed, are presented in Figure 2. An increase in the local temperature
90
characterizes the combustion zone, what moves towards the bed surface when increasing the
91
dense bed temperature . The combustion mainly takes place in the bed dense region when
92
dense bed temperature is higher than 800°C, what is consistent with previous observations
93
reported in the literature [Dounit (2001), Pré et al. (1998), Pré-Goubelle (1997)]. The local
94
temperature decreasing above the combustion zone is attributed to heat losses.
95
Pressure drop profiles.
96
The effect of the dense bed temperature, Tbed, on the normalized standard deviation of
97
σ the pressure drop fluctuations, Δ P , measured at different heights is presented in Figure 3.
98
The methane combustion occurs close to the bed surface when the dense bed temperature,
99
Tbed, reaches the critical value of 800 °C, and the resulting exploding of methane bubbles
100
reaching the bed surface induces an increase in the local pressure fluctuations. When
101
increasing the bed temperature above the critical temperature, Tbed, the local pressure
102
fluctuations decrease again. Such observations have ever been reported by Baskakov and
103
Makhorin (1975).
104
Reaction behaviour.
105
The freeboard region plays an important role in the combustion behaviour, especially
106
when the dense bed temperature is lower than the critical value of 800 °C. At a dense bed
107
temperature, Tbed, of 700 °C, the major part of the methane conversion is realized in the
108
freeboard region 400 mm above the dense bed surface, as underlined by the mole fraction of
109
the stable species (CH4, O2, CO2) measured along the reactor axis (Figure 4-a). The
110
progressive decreasing in the local methane mole fraction indicates that the natural gas
111
combustion process in the freeboard region of the reactor is a progressive process.
112
The bell shape of the CO mole fraction profile confirms the successive nature of the
113
methane combustion reaction with CO formation as an intermediate species (Figure 4-b). This
114
was already observed experimentally and modelled by several authors [Dryer and Glassman
115
(1973), Westbrook and Dryer (1981), Bradley et al. (1977)].
116
III.2: Effect of operating conditions.
117
The experiments are carried out at dense bed temperatures lower than the critical
118
temperature of 800 °C (Table 1). An increase in the excess air factor above 1.1 has little
119
influence on the process behaviour, both in the dense and dilute regions. The temperature and
120
the methane conversion profiles obtained at different excess air factors are very close to each
121
other (Dounit, 2001). On the other hand, an increase in the total mass of solids in the reactor
122
from 9 up to 15 kg has no significant effect on the reaction progress in the freeboard, since no
123
reaction occurs in the dense bed region at these low bed temperatures.
124
Superficial gas velocity effect.
125
Figures 5 and 6 present the temperature and the methane conversion measured along
126
the reactor axis at 700 °C. Three values of gas velocity are considered, respectively 2, 3 and 4
127
times Umf at 20 °C, corresponding to gas velocities ranging between 10.5 and 26 times U m f
128
at dense bed temperature. The combustion zone moves upwards from the bed level to the
129
outlet of the reactor when increasing the superficial gas velocity. This occurs because an
130
increase in the gas velocity leads to a decrease in the mean residence time of reactants in the
131
reactor, and also to an enhancement of the projection height of solid particles. These projected
132
particles contributes to the dilute phase temperature increasing and thus on the combustion.
133
Results obtained at bed temperatures lower than 750 °C were similar to those presented in this
134
section for 700 °C.
135
Mean particle diameter effect.
136
The effect of the mean particle diameter has been investigated between 100 and
137
550 µm (runs C1, C8 and C9 in Table 1). The superficial gas velocity was kept constant,
138
equivalent to twice the minimum fluidization velocity of particles 350 µm in diameter at 20
139
°C. Figure 7 presents the methane conversion as a function of the height at 700 °C for each
140
size of particles. Decreasing the mean particle diameter makes the combustion zone move
141
upwards to the reactor outlet, what is a consequence of the significant enhancement of two
142
parameters: the direct contact surface between gas and solid particles, and the particles hold-
143
up in the freeboard region, caused by an increase in the bubbles velocity and diameter at the
144
dense bed surface. At constant fluidizing velocity, decreasing the particle size leads to an
145
increase in the excess gas velocity ( U g − U m f ) what affects strongly the bubble size and
146
velocity. Similar behaviour was observed at the other dense bed temperatures tested in this
147
work.
148
IV: REACTOR MODEL.
149
The model for natural gas combustion process in the dense zone of the reactor is based
150
on the bubble assemblage model introduced by Kato and Wen (1969) what has been improved
151
to count for the thermal transfer that occurs during the combustion. It has ever been validated
152
at dense bed temperatures greater than 850 °C [Pré et al. (1998)]. The model is supplemented
153
with a freeboard modelling to count for the strong effects we experimentally evidenced of the
154
freeboard on the global bed combustion when the bed temperature Tbed is kept lower than the
155
critical temperature of 800 °C. The following assumptions are considered: I. The freeboard region is fed only by the gas contained in bubbles; the temperature and
156
composition of which are that of the bubble phase at the bed surface.
157 158
II. The projected particles originate from the bubble wakes. They are projected by
159
packets having initial velocities equal to the bubble rise velocity at the dense bed
160
surface. III. The particle projection by packets makes only a fraction of their total surface
161
accessible to gas. A contact efficiency factor is thus introduced.
162
IV. The entrained particle flux, the contact efficiency factor and the solid hold-up along
163
the freeboard fall exponentially from their values at the dense bed surface.
164
V. The gas phase and the particle phase flows are plug flows without and with back flow,
165
respectively.
166
VI. The freeboard region is subdivided into a number of elementary compartments.
167 168
IV.1: HEAT TRANSFER IN THE FREEBOARD.
169
The following mechanisms are counted in the modelling:
170
(4)
the gas-to-particle convection.
171
(5)
the gas-particle-suspension-to-reactor-wall conduction and radiation.
172
(6)
the particle-to-particle radiation.
173
As a first approximation, the radial heat transfer resistance is assumed to be located in a thin
174
film at the reactor wall.
175
The convective heat transfer coefficient between gas and solid particles is modelled
176
using the Ranz and Marshall’s correlation.
177
IV.1.1: Gas-particle suspension to reactor-wall transfer.
178
The model is similar to that proposed by Kunii and Levenspiel (1991) for the dense
179
region of the reactor. The heat flow exchanged between the gas-particle suspension and the
180
reactor wall is the sum of conductive and radiative contributions: q q c q re
181
(1)
182
Since the extinction coefficient ( p 1,5 f v p d p ) is much higher than unity (3,85 for a
183
system with 1 % solid hold-up and 350 µm mean particle diameter), the gas-solid suspension
184
is considered as a grey surface as well as the reactor-wall. Thus, the radiative contribution can
185
be written:
186
qre
Tp4 Tw4 1 1 1 w g
(2)
187
where εg is the gas-particles suspension emissivity, what depends both on the emissivity of
188
the gas phase and the local particles concentration. Its value is estimated using the correlation
189
reported in the appendix.
190
The conductive contribution is governed by the conduction through a gas film near the reactor
191
wall. Its thickness can be estimated as Lg= d p / 2 . This contribution is expressed, according to
192
Kunii and Levenspiel, as :
193
194
qc
kg Lg
T p T w
(3)
IV.1.2: Axial radiative heat transfer between different regions of the freeboard.
195
Various experimental measurements have shown that when the natural gas combustion
196
occurs partially or totally in the freeboard zone, the temperature and the solid hold-up varies
197
continuously along the reactor axis [Dounit et al. (2001), Dounit (2001)]. These gradients
198
generate an axial radiative heat transfer which can significantly affect the chemical reaction.
199
An approach similar to that proposed by Bueters et al. (1974) was used in this work.
200
The freeboard zone is discretized into N compartments separated by virtual planes. The
201
radiative heat flux absorbed by each slice i of the freeboard and coming from all other
202
freeboard slices can be expressed as:
q tot i gi E i 1 E i 1
203
j i 2 j 1
N
j i 2
204
j
gi
1 j 1
k i 1
gk
(4)
j 1 E j gi 1 gk k i 1
where Ej represents the radiative heat flow emitted in one direction by slice j:
Ej
205
206
E
1 gj 2 A T j4 2
(5)
IV.2: HYDRODYNAMICS OF THE SYSTEM.
207
Modelling of natural gas combustion in the freeboard region of a fluidized bed reactor
208
necessitates prior knowledge of the solid hold-up and the flux of projected particles along this
209
region. In Table 2 we present the correlations used.
210
(Table 2).
211
IV.3 : MODEL EQUATIONS.
212
The mass balance is written for every principal species present in the gas phase (CH4,
213
O2, N2, H2O, CO, CO2, and NO) while the heat balances are established separately both for
214
gaseous and particulate phases (Table 3).
215
(Table 3). The model equations are solved in three steps:
216 217
I.
Initially the heat and mass balances for each compartment of the dense region are solved,
218 219
II.
Secondly, for each slice of the freeboard zone, the heat and mass balances are solved.
220
III.
Finally the global heat and mass balances for the entire reactor is checked.
221
The systems of equations obtained both for the dense and for the freeboard zones are
222
non-linear algebraic systems solved using the Newton-Raphson method. More details can be
223
found in Pré-Goubelle (1997) and Dounit (2001). From this procedure, we get the axial
224
profiles of different species mole fractions present in the gas stream and both gas and particle
225
temperatures profiles along the reactor. The heat loss due to conduction and radiation to
226
external medium can also be computed.
227
V: DISCUSSION OF THE MODEL RESULTS.
228
V.1: KINETIC SCHEME.
229
The kinetic schemes available of Dryer and Glassman (1973), Westbrook and Dryer
230
(1981), and Bradley et al. (1977) have been tested and compared to our experiments. The one
231
of Dryer and Glassman (1973) gives the best predictions. Only the results obtained with this
232
kinetics scheme are discussed here:
233
3 CH 4 O2 CO 2 H 2 O 2
(R1)
234
1 CO O2 CO 2 2
(R2)
235
The rates of reactions R1 and R2 and the values of kinetic parameters involved are
236
reported in Table 4.
237
(Table 4).
238
At relatively low temperatures (less than 750 °C), good agreement has been obtained
239
for kinetic parameters giving the fastest reaction rate for methane transformation (n=13.4 and
240
E=197.768 kJ.mol-1), as shown in Figures 8 and 9. At dense bed temperatures closer to the
241
critical temperature of 800 °C, the best agreement was obtained for kinetic parameters of
242
methane transformation reaction closer to the mean values proposed by Dryer and Glassman
243
(n = 13 and E = 197.768 kJ.mol-1) as seen in Figures 10 and 11. Such modification in the
244
kinetics constant suggests that the mechanism of methane combustion does not occur only in
245
the homogeneous phase, and that sand particles may have a catalytic or inhibitive effect. This
246
is consistent with the previous experiments of Sotudeh-Gharebaagh (1998) showing a low
247
catalytic effect of sand particles at temperatures lower than 700°C, and an inhibitive effect at
248
temperatures between 750°C and 850°C.
249
(Figures 8, 9, 10 and 11).
250
V.2: EFFECT OF OPERATING PARAMETERS.
251
The excess air factor and the initial fixed bed height have a very weak effect on the
252
natural gas combustion process. The model matches well this observation (Dounit, 2001). The
253
strong effect of the gas superficial velocity and of the mean particle diameter is correctly
254
predicted by the model, as shown in Figure 12 at a dense bed temperature of 700 °C. In
255
addition, as shown in Figure 13, the measured mole fraction profiles are in good agreement
256
with the model predictions for different particles mean sizes.
257
(Figures 12 and 13).
258
VI: CONCLUSION.
259
In the present paper, a theoretical and experimental study of the influence of the
260
freeboard zone on the combustion of natural gas at temperatures lower than the critical
261
temperature have been presented. The experimental study demonstrates the progressive move
262
of the combustion zone towards the dense bed surface and thus into the interior of the dense
263
zone as the bed temperature rises. Increasing the fluidizing velocity or decreasing the mean
264
particle diameter induces the reaction zone displacement towards the reactor outlet. The
265
freeboard modelling combined with the dense bed modelling of Pré et al. (1998) predict
266
satisfactorily the reactor behaviour for different operating conditions. At dense bed
267
temperatures lower than the critical temperature, the methane burning is not an exploding
268
phenomenon mainly because of the existence of projected particles. The combustion occurs
269
progressively in the freeboard region and the conversion is total at the reactor outlet. Thus,
270
from a practical point of view, it seems necessary to use coarse particles with mean gas
271
velocities around 12 times the minimum fluidization velocity to ensure achievement of the
272
whole reaction in the reactor and to reach the maximum thermal efficiency of the reactor. The
273
fluidized bed reactor model gives useful information regarding the thermal efficiency of the
274
operation and permits estimation of the conditions at which the reactor efficiency is
275
maximum.
276
APPENDIX.
277
The gas-particle suspension emissivity g is computed according to the following correlation:
g p Cs g
278
(A1)
279
where g and p are respectively the emissivity of gas and particles. CS is a factor depending
280
on the gas and solid emissivity, mean diameter and concentration of solid particles.
281
The particles emissivity is computed, for coarse particles, according to the following
282
correlation:
283
p 1 exp 1,5 f v L d p
(A2)
284
where L is the characteristic length.
285
The gas emissivity is calculated using the emissivity of absorbent gas (CO2 and H2O) given
286
by the Hottel correlation :
287
H C CO CO C H O H O 2
2
2
2
(A3)
288
and the use of a blackening factor FE in order to take into account the presence of species in
289
the gas stream other than CO2 and H2O:
290
g FE 1 H FE
(A4)
291
In these correlations, CO2 and H 2O are the emissivity of carbon dioxide and water vapour,
292
respectively. C CO2 , C H 2O and are factors depending on the partial pressure of gas species,
293
the gas temperature and the characteristic length. The blackening factor for the gas stream is
294
equal to 1.2 according to Bueters and al (1974).
295
NOTATION.
296
A
: reactor cross sectional area (m2).
297
Cps
: specific heat of particles (J.kg-1.k-1).
298
Ci
: concentration of species i in gas stream (mol.m-3).
299
Dr
: reactor diameter (m).
300
dhi
: height of compartments i (m).
301
dp
: mean particle diameter (m)
302
Ei
: heat flux emitted by radiation from slice i and originate from this slice in one direction. The total heat flux emitted being 2.Ei (W).
303 304
F
: flux of entrained particles (kg.m-2.s-1).
305
F0
: flux of entrained particles at dense bed surface (kg.m-2.s-1).
306
fv
: volume fraction of particles in each freeboard compartment.
307
fw
: bubbles volume fraction occupied by the wake.
308
h
: height above the bed surface (m).
309
hgp
: heat transfer coefficient by convection between gas and solid particles (W.m-2.k-1).
310
hpw
: heat transfer coefficient by conduction and radiation between the gas-particle suspension and reactor-walls (W.m-2.k-1).
311 312
Hi
: molar enthalpy of species i (J.mol-1).
313
kg
: gas phase conductivity (W.m-1.K-1).
314
L
: characteristic height of a freeboard compartment (m).
315
Lg
: gas film thickness (m).
316
qc
: heat flux exchanged by conduction between the gas-solid suspension and reactor walls (w.m-2).
317 318 319
qre
: heat flux exchanged by radiation between the gas-solid suspension and reactor walls (w.m-2).
320
qtot i
: heat flux absorbed by slice i and coming from all other regions per unit surface (w.m-2).
321 322
ri
: reaction rate for every gas species i (mol.m-3.s-1).
323
R
: reactor radius (m).
324
Tbed
: dense bed average temperature (k).
325
Tg
: gas temperature (k).
326
Tp
: particle temperature (k).
327
Tw
: reactor wall temperature (k).
328
Ug
: superficial gas velocity (ms-1).
329
Umf
: minimum fluidising velocity of solid particles (m.s-1).
330
Ui
: interstitial gas velocity (m.s-1).
331
yi
: mole fraction of species i in gas stream.
332
Z
: height above the gas distributor (m).
333
Greek letters.
334 335
Pi
: Mean pressure drop (Pa).
336
mf
: dense bed voidage at minimum fluidising conditions.
337
: local freeboard voidage.
338
g
: gas-particles suspension emissivity.
339
w
: emissivity of reactor walls.
340
p
: particles emissivity
341
: gas-particle suspension density (kg/m3).
342
0
: gas-particle suspension density at bed surface (kg/m3).
343
s
: density of the particles (kg/m3).
344
: contact efficiency at height h.
345
bed
: contact efficiency at bed surface.
346
: Stefan-Boltzmann constant (5.67.10-8 w/m2.k4).
347
i
: Standard deviation of pressure drop fluctuations (Pa). Dimensionless numbers.
348
349
Reynolds number
Re
350
Prandtl number
Pr
[1]
g Cp g kg
Baskakov A.P., Makhorin K.E, (1975) "Combustion of natural gas in fluidized beds", Inst of Fuel symp. Series. 1 : Fluidized combustion, C3-1.
353 354
g
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[18] Sadilov P. V., Baskakov A. P., (1973) "Temperature fluctuations at the surface of a
394
fluidized bed with gas combustion occuring therein", Int. J. of Chem. Eng., 13 (3), pp
395
449.
396
[19]
397
reactor”, PhD Thesis report, Ecole Polytechnique de Montréal, QC, Canada.
398
[20]
399
fluidized bed" Ind. Eng. Chem. Res., 31, pp 999-1007.
400
[21] Wen, C.Y. and Chen, L.H., (1982) "Fluidized bed freeboard phenomena :
401 402 403
Sotudeh-Gharebaagh (1998) “Combustion of natural gas in a turbulent fluidized bed
Van Der Vaart D. R., (1992) "Mathematical modelling of methane combustion in a
entrainement and elutriation", AIChE J., 28 (1), pp117-128. [22] Westbrook C. K., Dryer F. L., (1981) "Simplified reaction mechanisms for the oxidation of hydrocarbon fuels in flames", Comb. Sci. and Tech. 27 (31).
404
[23] Yanata I., Makhorin K. E., Glukhomanyuk A. M., (1975) "Investigation and
405
modelling of the combustion of natural gas in a fluidized bed of inert heat carrier", Int.
406
J. Chem. Eng., 15 (1) pp 68-72.
407 408 409 410
[24] Yates J. G., Rowe P. N., (1977) "A model for chemical reaction in the freeboard region above a
fluidized bed", Trans. Instn. Chem. Engrs., 55, pp 137.
[25] Zenz F. A., Weil N. A., (1958) "A theoretical-empirical approach to the mechanism of particle entrainment from fluidized beds", AIChE J. 4 (4), pp 472-479.
Figure(s)
CH4
O2
CO2
4
CO
Flame detection
3
Cooling unit 2
(1)
5
2
2
Temperature and pressure drop acquisition devices
PID PID 4
Air 2
6 1 Reactor 2 manual valve. 3 Rotameter. 4 Electro-valve. 5 Pump. 6 Pneumatic valve
2
4
D.M
Figure 1: schematic of the set-up.
Natural gas
Figure(s)
Tbed = 650 °C Tbed = 800 °C
Tbed = 700 °C Tbed = 850 °C
Tbed = 750 °C
200 Dense bed
T-T
(°C)
100
0 0
20
40
60
80
100
120
140
-100
-200
Height (cm)
Figure 2: Evolution of the difference T-Tbed along the reactor for different dense bed temperatures.
Figure(s)
Z = 112 mm
Z = 250 mm
Z = 338 mm
Z = 438 mm
0,6 0,5
/P
0,4 0,3 0,2 0,1 0 600
650
700
750
800
850
900
Temperature (°C) Figure 3: Normalised standard deviation of pressure drop fluctuations at different heights in the reactor against dense bed temperature.
Figure(s)
CH4
20
CO2 O2
Mole fraction (%)
16 Freeboard
Dense bed
12 8 4 0 0
20
40
60
80
Height (cm) Figure 4-a: Mole fraction profiles of stable chemical species (Test C1, Tbed = 700 °C).
100
Figure(s)
CO mole fraction (%)
0,3
0,2
0,1
0 0
20
40
60
80
Height (cm) Figure 4-b: CO mole fraction profile (Test C1, Tbed = 700 °C).
100
Figure(s)
U = 2.Umf
U = 3.Umf
U = 4.Umf
Temperature (°C)
1000 900 800 700 600 500 0
20
40
60
80
100
120
Height (cm) Figure 5: Temperature profiles obtained at Tbed = 700 °C for three values of gas velocity.
140
Figure(s)
Ug = 2.Umf
Ug = 3.Umf
Ug = 4.Umf
Methane conversion rate (%)
100 80 60 40 20 0 0
20
40
60
80
Height (cm) Figure 6: Methane conversion rate profiles at Tbed = 700 °C for three values of gas velocity.
100
Figure(s)
dp = 100 µm
dp = 350 µm
dp = 550 µm
Methane conversion rate (%)
100 80 60 40 20 0 0
20
40
60
80
Height (cm) Figure 7: Methane conversion rate at 700 °C for three particle sizes.
100
Figure(s)
experimental measures Simulation: n=13,4 et E=47200 cal/mol
Temperature (°C)
900 800 700 600 500 0
20
40
60
80
100
120
140
Height (cm) Figure 8: Temperature profile of gas-particles suspension: Comparison between the experimental results and the model predictions (Tbed = 700 °C).
Figure(s)
Experimental measures : % CH4 Experimental measures : % CO2 Experimental measures : % O2 Simulation: n = 13,4 and E = 47200 cal/mol
Mole fraction (%)
20 16 12 8 4 0 0
20
40
60
80
100
Height (cm) Figure 9: Mole fraction profiles of the main species present in gas stream: Comparison between the experimental results and the model predictions (Tbed = 700 °C).
Figure(s)
Experimental measures Simulation : n = 13 and E = 47200 cal/mol
Temperature (°C)
900 800 700 600 500 0
20
40
60
80
100
120
140
Height (cm) Figure 10: Temperature profile of gas-particle suspension: Comparison between the experimental results and the model predictions (Tbed = 750 °C).
Figure(s)
Experimental measures : % CH4 Experimental measures : % CO2 Experimental measures : % O2 Simulation: n = 13 and E = 47200 cal/mol
Mole fraction (%)
20 16 12 8 4 0 0
20
40
60
80
100
Height (cm) Figure 11: Mole fraction profiles of the main species present in gas stream: Comparison between the experimental results and the model predictions (Tbed = 750 °C).
Figure(s)
Experiment al measures : U = 2Umf Experiment al measures : U = 3Umf Experiment al measures : U = 4Umf
Simulat ion : U = 2 Umf Simulat ion : U = 3 Umf Simulat ion : U = 4 Umf
Methane conversion rate (%)
100 80 60 40 20 0 0
20
40
60
80
100
120
Height (cm) Figure 12: Comparison between experimental results and model predictions: Effect of superficial gas velocity.
Figure(s)
Experimental measures : dp=100µm
Simulatio n : dp = 100µm
Experimental measures : dp=350µm
Simulatio n : dp = 350µm
Methane conversion rate (%)
100 80 60 40 20 0 0
20
40
60
80
100
120
140
Height (cm) Figure 13: Comparison between experimental results and model predictions: Effect of the mean particle diameter.
Table(s)
Experience
Solid mass (kg)
Ug/Umf at 20 °C
Excess air factor
dp (µm)
Variable parameter
C1
12
2
1,2
350
Reference
C2
9
C3
15
2
1,2
350
Solid mass
1,2
350
Superficial gas velocity
350
Excess air factor
100
Mean particle diameter
C4 C5 C6 C7 C8 C9
12
3 4
12
2
12
2
1 1,5 1,2
550
Table 1: Operating conditions of premixed air-natural gas combustion experiments.
Table(s)
Variable
Correlation
Referenc e
Solid hold-up
0 exp(a h)
[12]
Entrainment
F F0 exp(a h)
[13]
flux of particles projected at the bed surface Contact efficiency factor
1 f w 1 mf s U g U mf with fw = 0,25 2 1 1 bed exp a h with a 6.62 m-1
F0
Graphical correlation of Kunii and Levenspiel (1990) Table 2: Correlations used for the hydrodynamic parameters.
Exponential factor "a"
[13] [13] [12]
Table(s)
Mass balance
3 U g Ci
N
4 U g Ci
N 1
U g Ci
N 2
2 N ri dhN 0
i 1, Nc
gas phase N
U g NC yi H i 4 yi H i i 1 Tg i 1 N 6 1 N hgp R Tg T p 2 P dp
U g 3 Tg
NC
N 1
N
U g Tg
N
yi H i i 1 NC
N 2
dh N 0
particulate phase
Heat balance
F N 1 cp A T N 1 T F N 1 cp A T N 1 T s p ref d s p ref a 6 N N 1 N N hgp A Tg T p dh N qtot N d p N N 4 N Fa Fd cp s A T p Tref 2 A gN T p N h pw D r dh N T p Tw int
Particles to reactor wall heat transfer coefficient:
Tp Tw Tp2 Tw2 1 1 Lg 1 w g Gas-to-particle heat transfer coefficient: kg 1 1 2 0,6 Re 2 Pr 3 h gp d p h pw
kg
Table 3: Model equations (heat and mass balances).
i 1 2 3 4 5 0 6 7
CH4 O2 N2 CO2 H 2O CO NO
Table(s)
Reaction
Reaction rate expression
Kinetic parameters
n 13,2 0,2 R1
0,7 CO0,28 exp rCH 4 10 n CCH 4
rCO 10 CCO C n
R2
0, 25 O2
C
0, 5 H 2O
E 4,18 10 3 R T
E 4,18 4, 5 exp 10 R T
E 48400 1200 cal.mol -1
n 14,75 0,4
E 43000 2200 cal.mol -1
Table 4: Reaction rate expressions and kinetic parameters values according to Dryer and Glassman (1973).
